Process for 2-Ethylhexyl Acrylate Production Using Reactive Distillation

Article Cite This: Ind. Eng. Chem. Res. XXXX, XXX, XXX−XXX

pubs.acs.org/IECR

Process for 2‑Ethylhexyl Acrylate Production Using Reactive Distillation: Design, Control, and Economic Evaluation Mihai Daniel Moraru*,†,‡ and Costin Sorin Bildea‡ †

Department of Process Technology and Development, Hexion, Seattleweg 17, 3195 ND Pernis, The Netherlands Department of Chemical and Biochemical Engineering, University Politehnica of Bucharest, Str. Gh. Polizu 1-7, 011061 Bucharest, Romania

‡

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S Supporting Information *

ABSTRACT: 2-Ethylhexyl acrylate is industrially produced from acrylic acid and 2-ethylhexanol. The reaction, catalyzed by sulfuric acid, is performed in batch reactors followed by a continuous distillation section. Drawbacks such as corrosion, difficult product separation, and disposal of the spent acid can be avoided by using solid catalysts and process intensification technologies. In this paper, the design, control, and economics of a 20 kt/a reactive distillation process are presented. The column has a catalytic section (structured packing), containing 1600 kg catalyst Amberlyst 70, which converts the reactants. A stripping section of 4.1 m recovers high-purity acrylate, while a rectification section of 0.7 m achieves low acid concentration in the top vapors. After condensation, high-purity wastewater is obtained in the overhead decanter. The proposed control system shows robustness when capacity changes are required, or when the fresh reactants are contaminated. The economic analysis shows attractive economic potential and other key economic indicators.

1. INTRODUCTION 2-Ethylhexyl acrylate (2-EHA) is an important bulk chemical used in the production of homo and copolymers. Both these polymers find their end-use application in the manufacturing of printing inks, impregnating agents,1 coatings, paints, adhesives, binders for leather, paper and textiles,2 superabsorbents, thickeners, and plastic additives.3 The market has shown significant growth in the past few years, and this growth is anticipated to increase in the coming years.3 At the industrial scale, 2-EHA is produced by esterification of acrylic acid (AA) with 2-ethylhexanol (2-EH) in the presence of an organic solvent and a strong homogeneous catalyst as sulfuric acid.2 Reactors operated batchwise are employed to carry out the reaction (3−5 h of reaction time), followed by a distillation system for product recovery and purification.2 The use of strong homogeneous catalysts generates acidic, corrosive, and nonenvironmental friendly waste.3 These elements lead eventually to high maintenance costs and a continuously increasing difficulty to comply with environmental regulations. Using acidic solid-based catalysts eliminates corrosion, increases selectivity, and makes easier the acrylate separation and recovery.4 Recent publications have screened a large number of ion-exchange resins (solid catalysts) to perform and study this esterification reaction.3−7 In view of constructing new plants, making use of solid-based catalysts and applying process intensification technologies can reduce the cost of investment, footprint, and utility consumption, as well as minimizing the impact on environment. © XXXX American Chemical Society

In a recent article, we showed that production of 2-EHA using Amberlyst 70 as solid catalyst is feasible at industrial scale in conventional reaction−separation−recycle systems. The feasibility of reactive distillation processes is suggested in our previous abstract in a conference proceedings.8 A recent study7 reports batch experiments for 2-EHA production. Using Amberlyst 15, 80% reactant conversion is achievable by batch reactive distillation. The (steady-state) simulation of a large-scale continuous reactive distillation process presented in the same study7 predicts that 1000 kg/h of acrylic acid can be converted by using only 14 kg of catalyst placed on one reactive stage. The present paper continues the endeavor by developing the conceptual design, control and economic evaluation of a reactive distillation process. The analysis starts with the description of reaction kinetics and basic thermodynamics. A rigorous topology analysis of the vapor−liquid equilibrium diagram is described since it plays an important role in determining the technical feasibility of the process. After equipment sizing, the process dynamics is analyzed at various changes in operating conditions. The control structure is tested for increase and decrease of the fresh acid flow rate (i.e., capacity changes) and for contamination of the fresh reactants. Inferential one-point temperature control is applied to achieve Received: July 21, 2018 Revised: October 14, 2018 Accepted: October 24, 2018

A

DOI: 10.1021/acs.iecr.8b03368 Ind. Eng. Chem. Res. XXXX, XXX, XXX−XXX

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Industrial & Engineering Chemistry Research

al.,11 and density data are given by Rehberg et al.12 One set of binary interaction parameters is used to describe the vapor− liquid, liquid−liquid, and vapor−liquid−liquid equilibrium (VLE, LLE, and VLLE, respectively). Three pairs (water/AA, water/2-EH, and AA/2-EH) out of the six are available in the Aspen databanks. Details on the VLE of water/AA can be found in the recent paper of Niesbach et al.,13 and those of VLLE of water/2-EH are in that of Omota et al.14 The other three pairs (water/2-EHA, AA/2-EHA, and 2-EHA/2-EH) are estimated using the UNIFAC (uniquac functional-group activity coefficient) predictive model. The association parameters for the binary AA/water used in the equation of state are also available. Details on thermodynamics for this specific system can be found in our recent paper9 and the references therein. 3.2. Singular Points. Besides the information on the order of boiling point of pure components, the design of the reactive distillation-based process takes into account the azeotropy that this quaternary system presents. At 0.2 bar, the calculations find three minimum boiling heterogeneous azeotropes, which are also the lowest boilers in the system: two binary and one ternary azeotrope. Table 2 shows all seven singular points (three azeotropes and four pure components). One can remark that the azeotropes contain water in large amounts. The water/ 2-EH azeotrope (entry 2) was experimentally confirmed (at 1.013 bar) by Omota et al.14

the product purity, while the water purity is a result of the organic−aqueous liquid−liquid phase equilibrium. The study ends with a brief comparison between key economic indicators of the reactive distillation and reaction−separation−recycle processes. Aspen Plus v8.4 and v10.0 and Aspen Plus Dynamics v8.4 are used to aid the design and analysis of the process. Aspen Plus v10.0 is used only for the rate-based calculations presented in section 4.3.

2. REACTION KINETICS The equilibrium esterification reaction between 2-EH and AA with formation of water and 2-EHA is described by

and assumed to be the only reaction taking place in the system due to operation at relatively low temperatures and low residence times. This liquid phase reaction was studied by Komon et al.4 in the presence of Amberlyst 70, an acidic ionexchange resin acting as solid catalyst. The experimental data provided in ref 4 were used by Moraru and Bildea9 to determine the kinetic parameters of the pseudohomogeneous model described by r = k f (aacidaalcohol − (1/Keq)aestera water)

(2)

k f = k 0 exp(−EA /(RT ))

(3)

ln(1/Keq) = A + B /T

(4)

Table 2. Calculated Singular Points (Azeotropes and Pure Components) at 0.2 bara

where r [kmol/(kgcat·s)] is the reaction rate, kf [kmol/(kgcat·s)] is the forward reaction rate, ai (i = acid, alcohol, ester, water) are the liquid phase activities, and Keq [-] is the equilibrium constant. A [−] and B [K] are constants in the equilibrium constant equation. k0 [kmol/(kgcat·s)] is the pre-exponential factor in the Arrhenius equation, and EA [kJ/mol] is the activation energy. Table 1 shows the kinetic parameters. Table 1. Regressed Kinetic Parameters Using the Experimental Data from Komon et al.4,9

SP no.

typeb

TB [°C]

1 2 3 4 5 6 7

het. het. het. hom. hom. hom. hom.

59.7 59.7 59.8 60.1 95.3 134.8 159.8

water AA 2-EH [mol frac.] [mol frac.] [mol frac.] 0.9839 0.9848 0.9893 1

0.0121 0.0152

2-EHA [mol frac.] 0.0040 0.0107

1 1 1

a

parameter unit

k0 [kmol/(kgcat·s)]

EA [kJ/mol]

A [-]

B [K]

value

722.7

51.77

−8.5845

2438.5

Calculation is performed using the Distillation Synthesis tool in Aspen Plus b Singular point type: heterogeneous (het.) or homogeneous (hom.).

3.3. Ternary Diagrams. Figure 1 shows the four ternary diagrams for this quaternary system. These diagrams contain several equilibrium relationships, constructed solely using the Distillation Synthesis tool of Aspen Plus. The singular points (azeotropes and pure components) shown in Figure 1 follow the same numbering as that given in Table 2. The three ternary systems that contain water show large liquid−liquid immiscibility areas (calculated at 1.013 bar and 30 °C); the largest area is present in the ternary water/2EH/2-EHA since both the alcohol and the ester have a low mutual solubility with water. All ternary diagrams contain also the residue curves and distillation boundaries. One observation is that the distillation boundary defined by the singular points 3−1−2 falls into the immiscibility area and can be crossed by liquid−liquid phase separation. In the paper of Blagov and Hasse,15 the residue curve maps are called VLE diagrams; one remark is that in these diagrams, only the composition of the liquid phase (i.e., of the residue) is shown. For the system in this study, three of the four ternary systems (Figure 1) in which water is present show a liquid−

3. THERMODYNAMICS 3.1. Thermodynamic Method and Models. The UNIQHOC thermodynamic method in Aspen Plus and its default models are selected to calculate all physical properties necessary in process simulations. This thermodynamic method uses the UNIQUAC (universal quasi-chemical) activity coefficient model to describe the behavior of the liquid phase. The vapor phase is described using the Hayden− O’Connell (HOC) equation of state, including the chemical theory of dimerization to take into account the dimerization of acrylic acid. The model parameters for calculating the pure component properties, of all components in this system, are available in the Aspen databanks. One can consult the Aspen Plus product and easily evaluate the original sources based on which these parameters are determined. Details on acrylate vapor pressure are given by Stull,10 heat capacity data is given by Kulagina et B


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Figure 1. Ternary diagrams (mass-based): boiling liquid−liquid envelope calculated at 1.013 bar; residue curves maps (RCM) and singular points (SP) at 0.2 bar (SP numbering as in Table 2).

(2) Sketch the distillation boundaries of the ternary diagrams using the information determined in the previous step. (3) Identify the same singular points in Figure 9 of Kiva et al.16 and read their topological index (+1 or −1); the letters (a, c, d, h, and l), under the “identification” column in Table 3 (this paper), indicate the subfigures letters which are present in Figure 9 of reference by Kiva et al.16. (4) Establish for each point if it is a n+k , n−k , s+k , or s−k (i.e., determine their nomination) using the same Figure 9 of Kiva et al.16. (5) Sum-up the singular points that have identical nomination in order to determine the number of k-component nodes (N) and saddles (S), each having positive (+) or negative (−) index. (6) Check the rule of azeotropy given by eq 5

liquid boiling envelope very similar to the liquid−liquid envelope at 30 °C. Hence, these residue curve maps represent the VLLE diagrams of these particular systems. In the immiscibility regions, these residue curves are in fact the average composition of the two liquid phases. For the ternary system in which water is not present, the residue curves represent the VLE diagram. 3.4. Topological Analysis of the VLLE Diagram. On the basis of the analysis in section 3.3, one can conclude that this quaternary system presents vapor−liquid−liquid equilibrium. The topology of the VLLE diagram of this system is defined by its singular points.16 Therefore, the topological consistency of the diagram can be determined by checking the condition given by nc

k

∑ [2

(Nk+

+

Sk+

−

Nk−

−

Sk−)]

nc − 1

= ( −1)

2·(0 + 2 − 1 − 1) + 4·(0 + 0 − 0 − 2) + 8·(1 + 0 − 0 − 0) + 16·(0 + 0 − 0 − 0) = 0

+1 (5)

(6)

This condition, or the generalized rule of azeotropy, is frequently used in the literature for checking the consistency of vapor−liquid diagrams of nonreactive multi component systems.15−17 N+k , N−k , S+k , and S−k represent the number of k-component nodes (N) and saddles (S), each having positive (+) or negative (−) index; nc is the number of components in the system. The information on topological indices for quaternary systems given by Kiva et al.16 is used here to check the topological consistency. The procedure was previously described in detail,18 while here only a brief description is presented: (1) Find all singular points, type of the singular points and number of components forming the singular point.

This procedure is aided by the Azeotrope Search tool of Aspen Plus. Table 3 shows the results solving the steps 1−4, while the results of step 5 are given in Table 4. In step 6, using the values from Table 4, the condition (eq 5) is checked (eq 6). Since this condition is satisfied, the topology of the VLLE diagram for this nonreactive quaternary system is consistent.

k=1

4. PROCESS AND EQUIPMENT DESIGN 4.1. Feasible Operating Window and Process Synthesis. The feasibility region (pressure−temperature conditions) for reactive distillation is determined using thermodynamic considerations and reaction conditions. Once determined, the operating conditions of the column are selected within this feasible region where distillation and reaction C


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Industrial & Engineering Chemistry Research Table 3. Characterization of Singular Points of the Quaternary System Water/AA/2-EH/2-EHA SPa

type

number of components

identificationb

index

nominationc

1 2 3 4 5 6 7

UN S S S S S SN

3 2 2 1 1 1 1

l h h c d c a

+1 −1 −1 +1 −1 +1 −1

n+3 s−2 s−2 s+1 s−1 s+1 n+1

a

Numbering is the same as in Table 2 and Figure 1. bUsing Figure 9 from ref 16. cThe “nomination” term (in the procedure above and in this table) is adopted from ref 16; however, the symbols are different. For example, n+k in this study vs N+k in ref 16; this change is made in order to distinguish between the “nomination” term and the number of k-component singular points in eq 5.

Figure 2. Feasible operating window for using reactive distillation (area 3): overlap of distillation (area 1) and reaction (area 2) conditions. The T−P profile is for the equilibrium-based design; see section 4.2.2.

section. This arrangement leads to a counter-current flow of the reactants, which facilitates their conversion. Note that a different solution is adopted in the paper of Talnikar et al.,7 where the heavy reactant (2-EH) is fed at the bottom of the reactive section. 4.2. Equilibrium-Based Design. The equilibrium model considers that the column comprises of several theoretical stages. On each stage, the vapor and liquid phases are at thermodynamic equilibrium, while the reaction proceeds at a rate which depends on temperature, liquid phase reactants activity and amount of catalyst. 4.2.1. Process Description and Mass Balance. The design proposes a plant with a capacity of 20 000 t/a of 2-EHA with a purity of 99.5% mass. A recent study9 showed that a spacetime-yield of 1.57 kgproduct/kgcat/h in a conventional reactor− separation−recycle process using a catalytic fixed bed reactor is sufficient to achieve the capacity. This is equivalent to a catalyst amount of 1600 kg. Hence, the results presented hereafter are based considering this amount of catalyst and the equilibrium-based column design presented in section 4.2.2. The detailed process flow diagram is shown in Figure 3. Fresh 2-EH liquid is fed at the top of the catalytic bed (stage 3) of the reactive distillation (RD) column. Fresh AA is vaporized (EVAP) and fed below the catalytic bed (stage 17). The top vapors are condensed (COND) and phase separated in a decanter (DECANT). The organic phase (97.3% mass alcohol) of the decanter is returned as reflux to the column (stage 1), while the aqueous phase (99.9% mass water) is removed from the process. The material balance for selected streams is presented in Table 5. Figure 4 presents details on temperature, reaction rate, and concentration profiles along the column. The column operates in the interval 0.1−0.22 bar (top to bottoms). Temperature along the column (top-left diagram) varies between 91 and 162 °C, showing a smooth gradient from stage to stage. The rate of product formation (top-right diagram) is high on the stages 10−17 where the concentrations of AA and 2-EH in the liquid phase are high (bottom-left diagram); one observation is that the water concentration is very low, favoring 2-EHA

conditions overlap. Figure 2 presents the pressure−temperature diagram showing the distillation area (area 1) and reaction area (area 2), and the overlap between the two where the reactive distillation (area 3) is feasible. The boundaries delineating the distillation area are given by the vapor pressure of the lightest and heaviest components in the system. They are denoted by Pvap,light and Pvap,heavy, referring to the vapor pressure of 2-EH/water azeotrope and 2-EHA pure component, respectively. To be able to use cooling water of 25 °C in the condenser and high pressure steam of 42 bar in the reboiler, the temperature interval is limited between 45 and 240 °C. This creates a pressure operating interval between 0.001 and 1.8 bar. The reaction area is given by the temperature at which the catalyst starts to activate so the reaction can occur, and the maximum temperature the catalyst can withstand. The experiments of Komon et al.4 suggests a minimum temperature of 80 °C, while the maximum temperature that the catalyst withstands is 190 °C. Hence, the feasible reactive distillation operating conditions are 80− 190 °C and 0.001−1.8 bar. The conceptual layout of the process is developed based on the thermodynamic insights. Since the boiling points of the two reactants are in between the boiling points of the two products, the proposed column has a standard configuration: The middle section contains the catalyst and is the reactive distillation section where the reactants are converted. The top (rectifying) section obtains the light components as distillate vapors, while the bottom (stripping) section obtains the heavy component as bottom liquid. Therefore, the minimum-boiling heterogeneous azeotropes containing water (the lightest singular points) can be obtained as distillate vapors, while the heavy 2-EHA product can be obtained as bottoms liquid. When condensed, the azeotropes form a liquid−liquid heterogeneous mixture, and thus a decanter can be implemented to recover the organic phase and remove the aqueous phase from the process. The addition of feeds to the column is also standard: The alcohol is heavier than the acid and is fed as liquid at the top of the reactive distillation section, while the lighter acid is fed as vapor at the bottom of the

Table 4. Calculation of the Number of k-Component Nodes (N) and Saddles (S), with Positive (+) or Negative (−) Index k-component SP value

N+3 = ∑n+3 1

S−2 = ∑s−2 2

S+1 = ∑s+1 2 D

S−1 = ∑s−1 1

N−1 = ∑n−1 1


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Figure 3. Process flow diagram and plant-wide control of the reactive distillation process; the QC loop is specific to the CS1b structure only.

Table 5. Mass Balance of the Equilibrium-Based Design stream name

1

2

mole flow [kmol/h] mass flow [kg/h]

13.6 978

13.7 1779

3

4

5

5.9 663

13.6 245

13.7 2512

0.1538 0.0012 0.8440 997 ppm

0.9998 112 ppm 83 ppm trace

trace 234 ppm 0.0069 0.9928

0.0245 768 ppm 0.9731 0.0016

0.9990 447 ppm 596 ppm trace

trace 92 ppm 0.0049 0.9950

Mole Fraction water AA 2-EH 2-EHA

1 1 Mass Fraction

water AA 2-EH 2-EHA

1 1

as bottom liquid. The structured packing MellapakPlus 252Y is used for these two sections. For the reactive stages, Katapak-SP11 of Sulzer is employed. In this packing, the catalyst, with the bulk density ρcat,bulk = 770 kg/m3,occupies ϕcat ≈ 45% of the volume, being immobilized in wire gauze layers which are combined with layers of MellapakPlus.19 The separation efficiency is about 2 NTSM (m−1),19 equivalent to an HETP of 0.5 m. The number of reactive stages is determined as follows: (1) Select the number of reactive stages, and equally distribute the catalyst. (2) Run the simulation to find the column diameter (note that each simulation run makes use of a design specification which adjusts the bottom flow rate such that the required purity of the bottom product of 99.5% mass 2-EHA is achieved; the column diameter is determined using the Packing Sizing tool of the equilibrium-based calculations. (3) Calculate the maximum allowed amount of catalyst per stage; for the structured catalytic packing considered here, the maximum amount of catalyst (mcat,max) can be estimated using

formation with relatively high reaction rates. From stage 10 toward the top of the column, the reaction rate decreases due to decrease in AA concentration. The vapor composition profile (bottom-right diagram) shows an increase in water and 2-EH toward the top of the column, indicating the formation and stripping of the heterogeneous azeotrope out from the liquid phase. The operating conditions are well within the reactive distillation feasibility area showed in Figure 2. Operation at even higher reaction rates is possible by increasing the operating pressure to achieve higher temperatures. However, the reaction rate was determined in the temperature interval 80−120 °C. One observation here is that the temperature interval of the reactive zone in this design is 98−128 °C, only 8 °C outside the interval. 4.2.2. Column Design. This design is based on the results of the equilibrium-based simulation. The process employs 1600 kg of catalyst, which give the same space-time-yield as the tubular reactor used in a previous study.9 The column has 2 rectifying and 8 stripping stages, including the reboiler. These are sufficient to obtain the top vapors with the composition close to the heterogeneous azeotropes, and high-purity 2-EHA

π 2 ϕcatρcat,bulk mcat,max = HETP Dpacking 4 E

(7)


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Figure 4. Column profiles in the equilibrium-based design. Top left: temperature. Top right: reaction rate. Bottom left: liquid phase composition. Bottom right: vapor phase composition.

are used to design the column. The equilibrium-based simulation represents the starting point for setting up the rate-based simulation. Switching the calculations from one approach to the other, the rate-based modeling form of the RADFRAC block simulating the reactive distillation column is activated. Once the column internals have been specified (here, already available from the equilibrium-based simulation), Aspen makes an automatic selection for the correlations that calculate the mass and heat transfer coefficients. Note that multiple methods are available to calculate the transfer coefficients. In this particular case, the default selected correlation methods are Brf-85 and Chilton and Colburn (default naming), for calculating the mass and heat transfer coefficient, respectively. According to the Aspen Plus documentation, these methods are based on the papers of Bravo et al.21 and Taylor and Krishna,22 respectively. For consistency reasons, the interfacial area is calculated with the same method. The geometric characteristics of the packing, required by all these correlations, are taken by Aspen Plus from a confidential database provided by the vendor. Several preliminary runs indicated that despite being a light component the water is not removed in great extent from the liquid phase. As the reaction is reversible, the conversion of reactants is hindered. To increase the water vaporization necessary to achieve higher conversion, the number of stages is increased in the reactive zone, maintaining the same number of stages in the rectification and stripping zones. A brief sensitivity analysis (not presented here) shows that for the given amount of catalyst (1600 kg): (i) increasing the number of rectification stages leads to more acid in the top of the column and (ii) increasing/decreasing the stripping stages has little impact on the overall results (purities, duties).

(4) If the mass of catalyst is higher than the maximum, then increase the number of stages and reiterate; if not, then stop. Various design procedures are described in the literature; the above procedure follows to some extent the one outlined by Luyben and Yu.20 For 15 reactive stages, the catalyst loading is 106.6 kg/stage. The packing diameter, assumed equal to the column diameter, is Dpacking = 0.9 m. Thus, the maximum allowed amount of catalyst per stage is 108 kg (see eq 7), and the design is feasible. The total packing height is 12 m (= NT × HETP), resulting a packing volume of 7.6 m3. Increasing the number of reactive stages and maintaining the total amount of catalyst (i.e., 1600 kg), only slight changes in the mass and heat balance are observed. The same holds when the number of stages and the total amount of catalyst increase. 4.2.3. Equipment Sizing. The equipment sizing uses the material and energy balance generated by the Aspen Plus simulation. The major hold-ups required in the dynamic simulations and the main equipment sizes based on which the equipment cost is estimated are of interest. The diameter of the column is calculated by the packing sizing tool of the equilibrium-based calculations. The sump and the decanter are sized considering 10 and 20 min residence time, respectively. The heat exchangers (EVAP, COND, COOL) are designed based on a heat transfer coefficient of 930 W/(m2·K); in the dynamic simulations, the heat exchangers hold-up is neglected (i.e., an instantaneous model is used). The main dimensions are shown in the process flow diagram (Figure 3). 4.3. Rate-Based Design. In this section, the results of the rate-based simulation (i.e., considering the rate at which the mass and heat transfer takes place between the vapor and liquid phases, which are not in thermodynamic equilibrium) F


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from one phase to the other, in each of the column sections. The water and 2-EH have a positive net transfer rate (in the reactive-distillation section), indicating the transfer from the liquid to the vapor phase. This behavior is in agreement with the vapor (bottom-left diagram) and liquid composition profile (bottom-right diagram) which show an increase in water and 2-EH toward the top of the column. As mentioned, this indicates the stripping of the heterogeneous azeotrope out of liquid phase. The boil-up in the reboiler and the vapor feed at the bottom of the reactive distillation section provide vapors of 2-EHA and AA, respectively. These components have a negative transfer rate indicating the transfer from the vapor to the liquid phase; again, this behavior is in agreement with the vapor and liquid composition profiles of these two components. 4.4. Equilibrium versus Rate-Based Design. These two approaches provide two different designs. While the mass and energy balances are very similar (slight change on the water purity), the main dimensions of the reactive distillation column (Table 7) show some important differences. The equilibriumbased design gives a short and fat column, while the rate-based design gives a tall and thin column. The rate-based simulation reveals that removal of water from the liquid to the vapor phase is more difficult when compared to the prediction of the equilibrium model. Consequently, the liquid-phase water concentration is higher, and the rate of reaction decreases. Thus, a higher packing height is required in the reactive-distillation section, twice the value from equilibrium-based design. The total packing volume however is only about 22% larger.

The column design follows the same steps as proposed in section 4.2.2. The only difference is that the diameter is found using the Design tool provided by the rate-based calculation approach. During this iterative procedure, the number of stages is increased considerably; therefore, more and more space to accommodate the catalyst becomes available. This leads to the conclusion that a packing with a higher efficiency (i.e., lower HETP) and a lower catalyst volume fraction can be used to increase the mass transfer rates. Thus, Katapak-SP13 of Sulzer is selected. In this packing, the catalyst occupies ϕcat ≈ 20% of the volume.19 The separation efficiency is about 3 NTSM (m−1),19 equivalent to an HETP of 0.333 m. In rate-based calculations, this latter number is selected as the packed height per stage (H), while the HETP is back-calculated using H and the vapor and liquid composition profiles for all components after the simulation is converged. Applying this procedure gives a number of 49 reactive stages; the number of rectification and stripping stages remains the same as in the equilibrium-based design. The catalyst loading is 32.7 kg/stage. The packing diameter, assumed equal to the column diameter, is Dpacking = 0.71 m, allowing a maximum amount of catalyst per stage of 40.3 kg (see Equation 7). The total packing height is 24.5 m (= sum of back-calculated HETP of each stage), resulting in a volume of 9.7 m3. This packing can be distributed in 6 sections, each of about 4 m and provided with a liquid distributor on top of the section. The back-calculated HETP values vary between 0.335 and 0.564 m. Table 6 presents the mass balance of the rate-based calculations; the numbering of the streams remains the same as previously described. There are only minor differences compared with the mass balance of the equilibrium-based design. Some unconverted acrylic acid escapes in the top of the column and contaminates the water stream.

5. PLANT-WIDE CONTROL The objective of the control structure in this process is to keep the inventory and the product purity at the required set-points, despite any process changes. The dynamic behavior of the plant is studied for two types of changes: (i) increase/decrease of the plant throughput and (ii) contamination with water of the fresh reactants. The equilibrium-based simulation is used to study the process dynamics. 5.1. Control Structure and Controllers Parameters. Figure 3 presents the plant-wide control structure of the process. This specific structure has been previously employed18 in control of the n-butyl acrylate process. Two versions of this structure are studied, namely, the Control Structure 1a (CS1a) and Control Structure 1b (CS1b). The main difference is that CS1a uses a temperature control loop to maintain the purity of the product, while CS1b also brings a concentration control loop in cascade with the temperature controller. For the rest of the process, the same controllers having identical settings are used. The plant-wide control works as follows. The flow rate of fresh AA is fixed, setting in this manner the capacity of the plant. To satisfy the reaction stoichiometry, 2-EH is fed by using a feedback control loop, so none of the reactants accumulates or depletes. One mode to implement this plantwide control concept is to monitor the L/D ratio (i.e., organic/ aqueous). When 2-EH is fed in excess, the alcohol cannot leave the process with the bottom stream because the product purity is controlled by the bottom-temperature control loop. The excess can neither leave the process with the water stream from the overhead decanter due to the constraint imposed by the liquid−liquid equilibrium. Therefore, 2-EH starts to accumulate, and in consequence reflux L and reflux ratio L/D start to

Table 6. Mass Balance of the Rate-Based Design stream name

1

mole flow [kmol/h] mass flow [kg/h]

13.6 978

water AA 2-EH 2-EHA water AA 2-EH 2-EHA

1

1

2

3

13.6 5.9 1775 664 Mole Fraction 0.1607 0.0168 1 0.8044 0.0182 Mass Fraction 0.0258 0.0108 1 0.9336 0.0298

4

5

13.6 246

13.6 2507

0.9984 0.0016 90 ppm trace

trace 605 ppm 0.0067 0.9927

0.9932 0.0062 646 ppm 5 ppm

trace 237 ppm 0.0048 0.9950

Figure 5 presents details on temperature, reaction, and mass transfer rates, and concentration profiles along the packing height. The column operates in the interval 0.1−0.28 bar (top to bottoms). The rate-based model predicts a temperature interval between 85 and 169 °C (top-left diagram, left y-axis), slightly different compared with the one from the equilibriumbased column. The temperature interval of the reactive zone is 89−140 °C, which is about 20 °C outside the interval for which the reaction rate parameters were estimated. The specific reaction rate (kmol/h/m) is shown in the same graph (top-left diagram, right y-axis). The specific mass transfer rates (top-right diagram) show how each component moves G


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Figure 5. Column profiles in the rate-based approach. Top-left: temperature and specific reaction. Top-right: specific mass transfer rate. Bottomleft: liquid phase composition. Bottom-right: vapor phase composition.

manipulating the cooling water flow rate to the cooler (COOL). For the CS1b structure, a concentration controller (QC) in cascade with the temperature controller (TC) is implemented to control the product purity. The set-point of the concentration controller is 0.5% mass (sum of AA + 2-EH mass fraction), in order to obtain a product purity of 99.5% mass 2-EHA; the output of QC is the set-point to TC. The action of QC is direct: When the sum AA + 2-EH mass fraction increases, the temperature set-point is increased and more steam to the reboiler is added. As a consequence, the boil-up increases achieving in this way higher product purity. For this specific control structure (i.e., CS1b), two cases are presented with respect to the configuration of the concentration analyzer (A): (i) 5 min sampling interval and 5 min dead time and (ii) 30 min sampling time and 30 min dead time. The key controllers, namely, TC, RC, and QC, are tuned using the closed loop Auto Tune Variation (ATV) method available in Aspen Plus Dynamics. The tuning sequence is as follows: (1) tune TC using some parameters for RC, with the QC loop not closed; (2) tune RC using for TC the parameters just determined and with the QC loop not closed; (3) close the loop and tune QC using for TC and RC the parameters just determined. All controllers are tuned in the same way. The stability limit is determined by using a relay amplitude of 5% of the controller output range from which the ultimate gain (Ku) and ultimate period (Pu) are determined. Then, the gain (Kc) and integral time (Ti) are calculated using the Tyreus−Luyben tuning rules. Applying these steps for the first case (5 + 5 min, sample time + dead time), the values Kc = 0.0681%/% and Ti = 43.6

Table 7. Designs Comparison parameter mcat [kg] NT [-] Dpacking [m] Hpacking [m] rectification reactive-distillation stripping Vpacking [m3] Qreb [kW]

equilibrium-based

rate-based

1600 24 + 1 0.9 12 1 7.5 3.5 7.6 370

1600 59 + 1 0.71 24.5 0.7 19.7 4.1 9.7 381

increase. Monitoring L or L/D, it is possible to detect the accumulation or depletion of the alcohol and to act on the imbalance. This control philosophy is implemented using the RC feedback controller (see Figure 3), which adjusts the flow rate of fresh 2-EH such that L/D remains at the required value. In this way, the fresh 2-EH follows the fresh AA flow rate, maintaining its plant-wide inventory. At the equipment level, standard process control is implemented. The top-pressure in the column is controlled by manipulating the vapor flow rate; the temperature of the condensate is controlled by manipulating the flow rate of cooling water. The organic level in decanter is controlled by manipulating the reflux and the aqueous level by the flow rate of wastewater. The temperature of fresh AA is controlled by the steam to the evaporator (EVAP), the sump level by the bottom flow rate, and the temperature on stage 21 by the steam flow rate to the reboiler. Finally, the product is cooled by H


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Figure 6. Dynamic simulation results (flow rates and products purities) for throughput changes: +25% (top) and −25% (bottom); left: CS1a (no concentration controller), right: CS1b (with concentration controller).

Figure 7. Dynamic simulation results (flow rates and products purities) for contamination of fresh feeds: 5% mass water in fresh AA (top) and 5% mass water in fresh 2-EH (bottom); left: CS1a (no concentration controller), right: CS1b (with concentration controller).

I


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When the fresh alcohol is contaminated, CS1a is not able to maintain the required product purity (bottom-left diagram), the purity decreasing from 99.5 to 95.7% mass; the water purity is practically not affected. In the case of CS1b, the product purity is immediately forced back to its initial value after a very small decrease. In contrast with the previous case, in which the fresh acid is contaminated, the production capacity remains constant. The control structure adjusts (i.e., increases) the fresh alcohol flow rate to meet the reaction stoichiometry. Again, the comparison between the “fast” and “slow” analyzer shows that measuring the composition every 30 min can be used for controlling the acrylate purity (right diagrams).

min are obtained. However, the results of the dynamic simulations, presented in sections 5.2 and 5.3, prove that the parameters determined in the first case are conservative and can be used also for larger delays in concentration measurements. This means that measurement techniques based on chromatography that analyze a product sample every 30 min can be successfully used in control loops. All controllers tuning parameters are shown in the Supporting Information. 5.2. Change of Plant Throughput. The change of throughput is made by changing the fresh flow rate of AA by ±25% from its nominal value. At the start of the dynamic simulation, the process is in steady state. After 2 h, the flow rate is ramped to the new value during 1 h interval. Figure 6 presents the results. The top diagrams correspond to the +25% change of fresh AA, while the bottom diagrams to the −25% change. The left diagrams correspond to the CS1a control structure (i.e., without the concentration controller), while the right diagrams correspond to the CS1b structure (i.e., with the concentration controller). All diagrams show the same variables on their y-axis: selected flow rates on the primary yaxis, and purity of the product and water streams on the secondary y-axis (i.e., the concentration of the acrylate in the product stream, and the concentration of water in the water stream). For the +25% change, the CS1a is not able to maintain the product purity, the concentration of 2-EHA dropping from 99.5 to 97% mass. In contrast, the CS1b structure is able to maintain the required purity. Since there are no differences in controlling the top of the column, both control structures bring the water purity practically at the initial steady-state level. One observation is that for a relatively short period of time (roughly, 5 h), the water purity is lower (the lowest value reached is about 91% mass). Regarding the flow rates, they behave in a similar manner in both control cases. In the case of the CS1b control structure, two analyzers in the control loop are compared. A “fast” analyzer (solid lines: 5 + 5 min, sample + dead times) is compared with a “slow” analyzer (dashed lines: 30 + 30 min, sample + dead times). The differences in results are showed for the purity of the product stream (the right diagrams). These are not large. Thus, the conclusion is that analyzing the product composition every 30 min seems sufficiently frequent to be used as input data in a concentration control loop. 5.3. Contamination of Fresh Feeds. The changes are introduced in the same way as for the change in throughput. Figure 7 presents the results. The top diagrams correspond to the contamination with water of the fresh AA stream, while the bottom diagrams correspond to the fresh 2-EH stream contamination; in both cases, 5% mass water in the fresh streams is introduced. The results of the CS1a are on the left diagrams, while those of CS1b are on the right diagrams. Again, all diagrams have the same variables on their y-axis (flow rates on the primary y-axis, and purities on the secondary y-axis). When the fresh acid is contaminated, there are no major changes in the plant behavior, except a small decrease in the production capacity; this holds for both control structures. Since the flow rate of fresh acid is fixed (which sets the plant capacity) and becomes contaminated with water, the flow rate of AA fed to the process is lower. Therefore, 2-EHA is produced in a lower amount. The control structure adjusts (i.e., decreases) the flow rate of fresh alcohol, such that the reaction stoichiometry is met.

6. ECONOMIC EVALUATION 6.1. Calculation Basis. On the basis of the mass and energy balance, and the size of the major equipment items (resulted by applying both the equilibrium- and rate-based approach), a rough economic evaluation is made. In essence, any economic calculation is based on the cost of raw materials and product, utilities and cost of equipment. Several key economic indicators are calculated (see eqs 8−14): the economic potential (EP), the specific economic potential (EPspec), and the specific costs of reactants, utilities, amortization, and production. EP and its components are expressed in $/a, while the specific costs are all in $/ton of product. The components of EP are the cost of product (Cproduct) and reactants (Creactants), cost of utilities (steam, cooling water, and wastewater) and amortization cost of the total capital investment (CTCI). The amortization cost (Camortization) is the CTCI divided to a payback period of 3 years (eq 10). CTCI is the individual factored method of Guthrie, described in detail by Seider et al.,23 to which the costs for royalties and start-up are added (suggested by the same reference). CTCI for the column sizing made using the equilibrium-based design is described in detail in the Supporting Information. The CTCI for the rate-based design is very similar. The cost of equipment is based on the relations provided in the book of Dimian.24 EP = Cproduct − Creactants − Cutilities − Camortization

(8)

EPspec = EP/P2‐EHA

(9)

Camortization = C TCI/payback period

(10)

Cspec‐reactants = Creactants/P2‐EHA

(11)

Cspec‐utilities = Cutilities/P2‐EHA

(12)

Cspec‐amortization = Camortization /P2‐EHA

(13)

Cspec‐production = (Creactants + Cutilities + Camortization)/P2‐EHA (14)

The prices for the product, acid, and alcohol are taken from the MOLBASE (http://www.molbase.com): 2258 $/ton 2EHA, 1381 $/ton AA, and 1738 $/ton 2-EH were the reference prices in July 2017. Another large source that shares chemical prices, ICIS (https://www.icis.com), gives similar values for 2-EHA and 2-EH; no price for technical grade AA is provided. The specific cost of medium-pressure steam is 8.22 $/GJ (11 bar).25 The cost of cooling water is 0.354 $/GJ.26 The wastewater cost has two components: costs due to J


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Industrial & Engineering Chemistry Research Table 8. Key Economic Indicators and Process Parameters of 2-EHA Production Processes process

RSR-A

capacity EP EPspec Cspec‑reactants Cspec‑utilities Cspec‑amortization Cspec‑production

[t/a] [$/a] [$/ton] [$/ton] [$/ton] [$/ton] [$/ton]

XAA conversion X2‑EH conversion water purity

[%] [%] [%]

20095 8124761 404 1768 35 50 1854 99.97 99.29 99.8

RSR-B

RSR-C

Economic Indicators 19902 7584084 381 1786 39 52 1877 Process Parameters 99.02 99.29 96.1

concentration of organics and due to the total flow rate; namely, 500 $/ton of organics27 and 27 $/ton of wastewater,24 respectively. 6.2. Economic Indicators and Comparison with the RSR Processes. For the equilibrium-based design, the EP is estimated at roughly 9 286 000 $/a: the product brings 45 371 000 $/a, while the production costs are 36 085 000 $/a (reactants 35 531 000 $/a, utilities 187 000 $/a, and amortization 368 000 $/a). A breakdown of utility costs shows that 68.6% is spent on steam, 28.8% is spent on wastewater treatment, and only 2.6% is spent on cooling water. Regarding the amortization, the shares are as follows: column of 56.4%, heat exchangers of 38.4%, and the decanter of 6.2% only. The other key economic indicators and some important process parameters are presented in Table 8. From an economic perspective, the differences in results generated using the equilibrium- and rate-based calculations are very small. Besides the results of this study on the reactive distillation process, the results of a recently published study9 on three conventional reaction−separation−recycle (RSR) processes are also included. All plants produce the same acrylate and have the same capacity; for a fair comparison, all prices, equations and assumptions that influence the economic analysis are the same as reported in the previous study.9 On the basis of these considerations, the following comparisons are made: (1) In comparison with the RSR processes, the RD process shows better economic indicators and similar process parameters. (2) The EP (or EPspec) is about 12% higher due to lower specific cost of utilities (Cspec‑utilities) and specific cost of amortization (Cspec‑amortization). (3) Although on the same order of magnitude, Cspec‑utilities for the RD process is a factor of 3.5 lower than that of the RSR-A process; Cspec‑amortization is a factor of 2.5 lower (rate-based design).

19937 7048299 354 1773 53 78 1904

RDe

RDr

20094 9286127 462 1768 9.3 18 1796

20057 9157127 457 1771 9.6 20 1801

99.18 99.29 96.7

99.97 99.3 99.9

99.77 99.31 99.3

mass purity acrylate and 99.9% mass purity wastewater. The rate-based design proposes a column with a much longer packing section in the reactive distillation zone (19.7 m), and similar rectification (0.7 m) and stripping (4.1 m) sections. For this design, Katapak-SP13 is used. The dynamic simulations using the equilibrium-based design show that the control structure is able to achieve large throughput changes, and to cope with the contamination with water of both fresh reactants. However, one concentration measurement is necessary to achieve tight control of the product purity. The economic potential is raw materials-driven. The net earnings are 460 $/ton of acrylate at a production cost of 1800 $/ton. Compared to the reaction−separation−recycle process, the utility and amortization costs of the reactive distillation process are lower with a factor of 3.5 and 2.5, respectively.

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ASSOCIATED CONTENT

* Supporting Information S

The Supporting Information is available free of charge on the ACS Publications website at DOI: 10.1021/acs.iecr.8b03368.

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Calculation of the total capital investment (CTCI) from the total installed cost (CTBM) of process equipment; detailed controllers parameters (PDF)

AUTHOR INFORMATION

Corresponding Author

*E-mail: [email protected]. ORCID

Mihai Daniel Moraru: 0000-0001-9223-9913 Costin Sorin Bildea: 0000-0001-7707-1366 Notes

The authors declare no competing financial interest.

7. CONCLUSIONS 2-Ethylhexyl acrylate production using the reactive distillation technology and solid-based catalysts is feasible; 20 000 ton of acrylate per annum can be achieved in a catalytic column− overhead decanter system. The equilibrium-based design model indicates that a reactive distillation section of 7.5 m containing Katapak-SP11 (∼1600 kg catalyst Amberlyst 70) is able to convert the reactants and achieve the required capacity. A rectification section of 1.0 m and a stripping section of 3.5 m, both containing MellapakPlus, are sufficient to achieve low concentrations of acrylic acid in the overhead vapors and highpurity 2-ethylhexyl acrylate at the bottom. This process achieves complete conversion of reactants, and obtains 99.5%

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ACKNOWLEDGMENTS C.S.B. gratefully acknowledges the financial support of the European Commission through the European Regional Development Fund and of the Romanian state budget, under the grant agreement 155/25.11.2016 (Project POC P-37-449, acronym ASPiRE).

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REFERENCES

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Process for 2-Ethylhexyl Acrylate Production Using Reactive Distillation

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